JP2008506004A - Hydrogenation of aromatic compounds and olefins using mesoporous catalysts - Google Patents

Abstract

A method of hydrogenating a hydrocarbon feedstock containing unsaturated components, having at least 97 volume percent interconnected mesopores based on mesopores and micropores, and at least 300 m ² / g Providing a catalyst comprising at least one Group VIII metal on an amorphous mesoporous inorganic oxide support having a BET surface area and a pore volume of at least 0.3 cm ³ / g; and the presence of said catalyst And contacting the hydrocarbon feedstock with hydrogen under hydrogenation reaction conditions.
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Description

  This application is a continuation-in-part of co-pending US patent application Ser. No. 10 / 886,993, filed Jul. 8, 2004, incorporated herein by reference.

  The present invention relates to a process and catalyst for hydrogenating aromatics and olefins in hydrocarbon streams, preferably but not limited to hydrocarbon distillates.

  Removing aromatics from various hydrocarbon distillates (eg, jet fuel, diesel fuel, lube base stock, etc.) is because monocyclic and polycyclic aromatic compounds can be mixed in diverse ways. Have difficulty. In most refiners, dearomatization requires significant capital investment, but with the benefits associated therewith. The aromatics content of the distillate is closely related to the cetane number, which is the main quality standard for diesel fuel. The cetane number is highly dependent on the paraffin ratio, the saturation rate of the hydrocarbon molecule, and whether the hydrocarbon molecule is a linear molecule or has an alkyl side chain attached to the ring. Distillate streams that are mostly aromatic molecules with little or no alkyl side chains generally have low cetane quality. On the other hand, distillate logistics having a high paraffin ratio generally have high cetane quality. Jet fuel quality also depends on whether the aromatic content is lower due to the aromatic / smoke point relationship. Most jet fuels are limited by aromatics content to a maximum of 25 volume percent by standard.

  The increasing demand for distillates with higher paraffin content is also the impact of the regulatory environment. Dearomatization is becoming increasingly important due to government bills that require a substantial reduction in the content of distillate aromatics and polycyclic aromatics. Current US Environmental Protection Agency standards for diesel fuel limit the aromatic content of diesel fuel to a maximum of 35 volume percent. California's diesel fuel specification is up to 10 volume percent.

  In many parts of the world, there is a phenomenon called “dieselization”. This represents an upward shift in the demand ratio of diesel / gasoline fuel as the demand for fuel generally increases. Global demand for diesel fuel is expected to double between 2000 and 2010, partly due to economic growth, due to efforts to combat global warming and even from general demand for fuel efficiency Is done. One approach to meeting these requirements would be to shift the use of lower quality home heating oil to automotive diesel fuel. This is thought to increase the need for desulfurization and dearomatization.

  However, higher distillate paraffin content results in more severe reaction conditions with conventional metal hydride catalysts such as cobalt, molybdenum, nickel, and tungsten. In recent years, it has been proven that dearomatization catalysts having high activity can be obtained by using mixed noble metals on a support or zeolite.

  US Pat. No. 5,151,172 to Kukes et al. Discloses a process for hydrogenating a distillate feedstock with a catalyst comprising a combination of palladium and platinum on a zeolite (ie, mordenite) support.

  U.S. Pat. No. 5,147,526 to Kukes et al. Discloses a distillate with a catalyst comprising a combination of palladium and platinum on a Y-type zeolite support containing about 1.5 wt% to about 8.0 wt% sodium. A method for hydrogenating raw materials is disclosed.

  US Pat. No. 5,346,612 to Kukes et al. Discloses a process using a combination of palladium and platinum on a β-type zeolite support.

US Pat. No. 5,541,312 to Apelian et al. Discloses platinum and palladium on MCM-41, a mesoporous crystal support. The use of a mesoporous carrier has the advantage that mass transfer restrictions are suppressed because the pore system is significantly larger. However, although mesoporous carriers have better molecular access compared to zeolites, crystalline mesoporous materials are limited because the pores are not interconnected. Furthermore, the oxide variations that can be used in crystalline mesoporous supports without destroying the crystal structure of the support are very limited.
US Pat. No. 5,151,172 US Pat. No. 5,147,526 US Pat. No. 5,346,612 US Pat. No. 5,541,312

  Therefore, there is a need for a catalyst system that has mesopores that are highly connected to each other and that can select pore diameters from a wide range, and that can select inorganic oxide components in the structure more flexibly. Yes.

In this application, the method of hydrogenating the hydrocarbon raw material containing an unsaturated component is provided. The method has at least 97 volume percent interconnected mesopores, based on mesopores and micropores, and has a BET surface area of at least 300 m ² / g and pores of at least 0.3 cm ³ / g Preparing a catalyst containing at least one Group VIII metal on an amorphous mesoporous inorganic oxide support having a volume, and hydrogenating the hydrocarbon feedstock under hydrogenation reaction conditions in the presence of the catalyst. And contacting with.

  The present invention provides a mesoporous catalyst system that has mesopores that are highly connected to each other, that can adjust the pore diameter from a wide range, and that can select the inorganic oxide component in the structure more flexibly. To do. Furthermore, the system of the present invention allows the zeolite to be dispersed within the mesoporous matrix, thereby significantly improving access to the small crystal zeolite.

  The present invention saturates distillate hydrocarbon feed containing aromatic compounds and / or olefins using a catalyst containing one or more noble metals on the catalyst support to reduce unsaturated components in the feed. To provide a method.

  While other petroleum streams may also benefit from the present invention, the preferred distillate hydrocarbon feed processed in the present invention is about 150 ° F. (66 ° C.) to about 700 ° F. (371 ° C.), preferably Include any purified stream boiling in the range of 300 ° F. (149 ° C.) to about 700 ° F. (371 ° C.), more preferably about 350 ° F. (177 ° C.) to about 700 ° F. (371 ° C.). It is done.

  A feature of the present invention is that it has an ability to treat hydrocarbon raw materials having an aromatic compound content exceeding 20 mass%, 50 mass%, 70 mass%, or even 80 mass%.

  Distillate hydrocarbon feedstocks include high sulfur and low sulfur virgin distillates derived from high sulfur and low sulfur crude oils, coke distillates, catalytic cracking light and heavy catalytic cycle oils, bisbreaker distillates, and hydrogen Distillate boiling range products from additive crackers, FCC or TCC feed hydrotreaters, and residual hydrotreating equipment are included. In general, light and heavy catalyst cycle oil is a raw material component having the highest aromaticity (FIA) with an aromatic compound as high as 80% by mass (FIA). Most of the aromatic compounds contained in the cycle oil exist as monocyclic aromatic compounds and bicyclic aromatic compounds, and some exist as tricyclic aromatic compounds.

  Virgin stocks such as, for example, high sulfur and low sulfur virgin distillates have a lower aromatic content of 20% by weight (FIA). Generally, the content of aromatic compounds in the mixed hydrogenation equipment feedstock is from about 5% to about 80%, more typically from about 10% to about 70%, most typically about 20%. % To about 60% by weight. In many cases, the catalytic process proceeds so as to equilibrate the aromatic compound concentration of the product at a sufficiently low space velocity. It is more advantageous to do so.

  The sulfur concentration of the distillate hydrocarbon feed is generally a function of a mixture of high and low sulfur crude oil, hydroprocessing capacity per barrel of refinery crude capacity, and alternative treatment of distillate feed components. is there. The higher sulfur distillate feedstock components are generally coke distillates, bisbreaker distillates, and catalytic cycle oils. These distillate raw material components may have a total nitrogen concentration as high as 2000 ppm, but are generally from about 5 ppm to about 900 ppm.

  Particularly preferred raw materials in the present invention are hydrocarbon fractions in jet fuel and diesel fuel having a boiling point in the range of 150 to 400 ° C. Typical aromatic compounds contained in the raw material include monocyclic aromatic compounds, bicyclic aromatic compounds, and tricyclic aromatic compounds, particularly those that normally boil below about 343 ° C. Is mentioned. Examples of aromatic compounds contained in the raw materials include monocyclic aromatic compounds such as alkylbenzene, indane / tetralin, and dinaphthenebenzene, such as bicyclic aromatic compounds such as naphthalene, biphenyl, and fluorene. And tricyclic aromatic compounds such as phenanthrene and naphthenanthrene. Raw materials containing a significant proportion of polycyclic aromatic compounds are preferred (ie, up to 100 weight percent of the total aromatic compounds in such raw materials can consist of polycyclic aromatic compounds), usually The raw material to be treated in the invention contains a considerable proportion of monocyclic aromatic compounds, and the proportion of polycyclic aromatic compounds is relatively small. The content of the monocyclic aromatic compound in the raw material wholly aromatic compound is usually more than 50 mass percent. In this application, a typical hydrocarbon distillate or mixture thereof comprises at least about 10 volume percent aromatic hydrocarbon compound. Most preferred raw materials in the present invention are at least 10 volume percent, often at least 20 volume percent, usually more than 30 volume percent, generally in the range of about 10 to about 80 volume percent, often about 20 to 50 volume percent. A diesel fuel feedstock containing a wide range of aromatic-containing compounds. The greatest benefit of the method of the present invention is achieved when a high concentration of aromatic compound in the feed is saturated without substantial decomposition of the homocyclic aromatic compound.

  Other preferred feedstocks include hydrocarbons of lubricant viscosity. Treatment can be improved by using mineral oil lubricants or synthetic hydrocarbon lubricants. Examples of synthetic hydrocarbon lubricants include poly α-olefin (“PAO”) materials, ie conventional types of PAO using Friedel-Crafts type catalysts, and reduced Group VIB (Cr, Mo, W ) Both HVI-PAO materials produced using metal oxide catalysts.

  Mineral oil lubricants are generally characterized as having a minimum boiling point of at least 650 ° F. (343 ° C.). They can also be neutral, i.e. distillates, stocks with a 95% boiling point of 1050 ° F (566 ° C) or less, but residual lubricant stocks such as bright stocks are the same catalytic process. Can be processed by. This kind of mineral oil stock has historically distilled crude oil of suitable composition at atmospheric pressure and vacuum, then using solvents such as phenol, furfural, or N, N-dimethylformamide ("DMF") Produced by conventional purification methods that involve removing undesirable aromatic components by solvent extraction. Dewaxing to obtain the desired product pour point can be performed by solvent dewaxing or catalytic dewaxing techniques (or combinations thereof). In order to saturate the lube oil boiling range olefin that may be produced during the catalytic dewaxing step, it is particularly preferred that the hydrotreatment according to the present invention be carried out after any catalytic dewaxing treatment.

  Mineral oil stocks can also be produced by catalytic hydrocracking. Here, the unconverted high-boiling hydrocarbon stream acts as a waxy lubricating oil base. Following the hydrocracking process, the lubricating oil stock is dewaxed and hydrorefined to adjust fluidity and reduce olefins and possibly aromatics. This process, commonly referred to as “lubricating oil hydrocracking”, is often used when raw materials are insufficient with conventional lubricating oil processing methods or when high VI lubricating oil products are required. .

  The method is also applicable to the hydroprocessing of synthetic lubricating oils, especially poly alpha-olefins ("PAO") containing HVI-PAO type materials. Such types of lubricating oils are conventionally, for example, Friedel-Crafts type such as aluminum trichloride, boron trifluoride or boron trifluoride complex in the presence of water, lower alkanols or esters. It can be produced by polymerization or oligomerization using a catalyst. HVI-PAO type oligomers can be prepared using reduced Group VIB metal oxide catalysts, usually chromium on silica, by the methods described in US Pat. No. 4,827,064 or US Pat. it can. HVI-PAO materials include higher molecular weights made with lower oligomerization temperatures as disclosed in US Pat. No. 5,012,020. HVI-PAO materials are characterized by a branching ratio of less than 0.19 due to the use of a special reduced metal oxide catalyst during the oligomerization process.

  The lubricant material is hydrotreated in the presence of a catalyst comprising the mesoporous material of the present invention and optionally a metal component for hydrogenation with a binder.

  The hydrogenation reaction is conducted under conventional conditions at a temperature in the range of about 100 to about 700 ° F., preferably 150 to 500 ° F. Hydrogen is preferably under superatmospheric conditions, and the hydrogen partial pressure can vary up to about 2500 psi, but is usually in the range of about 690-10343 kPa (100-1500 psi). The hydrogen cycling ratio is generally taken as a stoichiometric excess in the range of 200% to 5000%, exceeding the amount required stoichiometrically for complete saturation. In order to maximize the purity of hydrogen, once-through circulation is preferred. The space velocity is generally in the range of 0.1 to 10 LHSV, usually 1 to 3 LHSV. The product of the hydrogenation reaction has a degree of unsaturation that is low enough to be consistent with the hydrotreatment. In most cases, hydrocarbon lubricant feedstocks having a bromine number greater than 5 are processed according to the method of the present invention resulting in a product having a bromine number of less than 3 and often less than 1.

  When a specific hydrotreatment facility is a two-stage treatment, it is often configured to perform desulfurization and denitrogenation in the first stage and to dearomatize in the second stage. In these operations, the feedstock entering the dearomatization stage has a substantially lower nitrogen and sulfur content and may have a lower aromatic content compared to the feedstock entering the hydrotreating facility. .

  The hydrogenation process of the present invention generally begins with a distillate feedstock preheating step. Prior to entering the furnace, the feed is heated in the feed / effluent heat exchanger to eventually reach the target reaction zone inlet temperature. The feed may be contacted with the hydrogen stream before, during and / or after preheating. The hydrogen-containing stream may be added to the hydrogenation reaction zone of the one-stage hydrogenation process or may be added to the first or second stage of the two-stage hydrogenation process.

  The hydrogen stream may be pure hydrogen or a mixture with diluents such as hydrocarbons, carbon monoxide, carbon dioxide, nitrogen, water, sulfur compounds, and the like. The purity of the hydrogen stream should be at least about 50% by volume hydrogen, preferably at least about 65% by volume, and more preferably at least about 75% by volume for best results. Hydrogen can be supplied by a hydrogen plant, catalytic reforming equipment, or other hydrogen production process.

The reaction zone may consist of one or more fixed bed reactors containing the same or different catalysts. The two-stage process can be designed to have at least one fixed bed reactor for desulfurization and denitrogenation and at least one fixed bed reactor for dearomatization. Fixed bed reactors often have multiple catalyst layers. Optionally, one fixed bed effluent can be cooled before being directed to a subsequent fixed bed. Multiple catalyst layers in one fixed bed reactor may contain the same or different catalysts. If the catalysts in the multi-layer fixed bed reactor are different, the first layer (s) are usually for desulfurization and denitrogenation and the subsequent layers are for dearomatization. When a multi-reactor system is used, the gas between the reactors undergoes heat “volatilization” to remove H ₂ S and NH ₃ . The product gas in these first stages causes reaction suppression and more seriously contaminates the noble metal (s) on the dearomatization catalyst.

  Since the hydrogenation reaction is usually exothermic, intermediate cooling via hydrogen injection may be performed. Other methods including intermediate heat transfer may be used. The two-stage method can reduce the temperature of the exotherm for each reactor shell and can better control the temperature of the entire reactor. This is important for safety and optimal catalyst efficiency and lifetime.

The reaction zone effluent is generally cooled and directed to a separator to remove hydrogen. One example is an amine washer. H ₂ S is sent to the sulfur recovery unit, and in many cases, NH ₃ is recovered as a purification byproduct. Part of the recovered hydrogen is recycled to the process and part is sent to other less burdensome hydrogenation units (eg naphtha pretreatment equipment) or discharged to external systems such as plants or refined fuel Is done. The hydrogen discharge rate is often controlled to maintain a minimum hydrogen purity and to remove hydrogen sulfide. The recycled hydrogen is generally compressed and refilled with “make-up” hydrogen and reinjected into the process for further hydrogenation. One preferred disposal strategy for low purity hydrogen is to return to the hydrogen plant loop, where much of the hydrogen is recovered upstream of the hydrogen unit by the absorber.

  The liquid effluent of the separation device is then processed in the cleaning device. Here, light hydrocarbons are removed and directed to a more appropriate hydrocarbon pool. The scrubber liquid effluent product is then generally transferred to a blending facility to produce a finished distillate product.

  The operating conditions in the hydroprocessing step of the present invention are from about 300 ° F. (150 ° C.) to about 750 ° F. (400 ° C.), preferably from about 500 ° F. (260 ° C.) to about best for best results. It includes an average temperature of the reaction zone of 650 ° F. (343 ° C.), most preferably from about 525 ° F. (275 ° C.) to about 625 ° F. (330 ° C.). At reaction temperatures below these ranges, the efficiency of hydrogenation decreases. If the temperature is too high, the thermodynamic limits of aromatic reduction may be reached, and non-selective hydrocracking, catalyst deactivation, and increased energy costs may occur.

  The method of the present invention generally provides from about 1379 kPa to about 13790 kPa (about 200 psi to about 2000 psi), more preferably from about 3448 kPa to about 10343 kPa (about 500 psi to about 1500 psi), and most preferably from about 4137 kPa to obtain the best results. It operates at a partial hydrogen pressure in the reaction zone ranging from about 600 psi to about 1200 psi. The hydrogen cycling ratio generally ranges from about 500 SCF / Bbl to about 20000 SCF / Bbl, preferably from about 2000 SCF / Bbl to about 15000 SCF / Bbl, most preferably from about 3000 SCF / Bbl to about 13000 SCF / Bbl for best results. Across. Reaction pressures and hydrogen circulation ratios below these can result in increased catalyst deactivation rates and reduced desulfurization, denitrogenation, and dearomatization efficiency. If the reaction pressure is too high, energy and equipment costs increase and margins are reduced.

The method of the present invention is generally about 0.2 hours- ¹ to about 10.0 hours- ¹ , preferably about 0.5 hours- ¹ to about 3.0 hours- ¹ , for best results. Preferably, it operates at a liquid space velocity of about 1.0 hour- ¹ to about 2.0 hour- ¹ . If the space velocity is too high, the hydrogenation can be reduced overall.

The catalyst carrier denoted TUD-1 is a three-dimensional amorphous, mesoporous inorganic oxide material, and at least 97 volume percent of meso-connected mesopores based on the micropores and mesopores of the organic oxide material. Contains pores (ie, less than 3 volume percent micropores), and generally contains at least 98 volume percent mesopores. Preferred methods for producing porous catalyst supports are described in US Pat. No. 6,358,486 and US patent application Ser. No. 10/764797 filed Jan. 26, 2004 (“Method For Making Mesoporous Combined Mesoporous and Microporous Inorganics”). Both of which are incorporated herein by reference. The average mesopore diameter of the preferred catalysts as measured by N ₂ porosimetry ranges from about 2nm~ about 25 nm. Generally, a mesoporous inorganic oxide is produced by heating a mixture of (1) a precursor of an inorganic oxide in water and (2) an organic pore forming agent at a predetermined temperature for a predetermined time.

  The starting material is generally an amorphous material, which can be, for example, one or more inorganic oxides such as silicon oxide or aluminum oxide, with or without additional metal oxides. . Silicon atoms can be partially substituted with metal atoms such as aluminum, titanium, vanadium, zirconium, gallium, manganese, zinc, chromium, molybdenum, nickel, cobalt, and iron, and the like. The inorganic oxide is preferably selected from the group consisting of silica, alumina, silica-alumina, titania, zirconia, magnesia, and combinations thereof. Additional metal may optionally be incorporated into the material prior to initiating the process of producing the structure containing mesopores. Also, after material preparation, the cations in the system may optionally be replaced with other ions, such as ions of alkali metals (eg, sodium, potassium, lithium, etc.). Alkali cations can determine any acidity remaining in the interior of TUD-1, especially in the case of Al-TUD-1 or Al-Si-TUD-1. Residual acidity causes undesirable decomposition reactions, thereby reducing the total production of liquid products.

Mesoporous catalyst supports are amorphous materials (ie, no crystallinity is observed with currently available X-ray diffraction techniques). The d interval of the mesopores is preferably about 3 nm to about 30 nm. Surface area of the catalyst support measured by BET _{(N 2)} is at least about 300 meters ² / g, preferably ranging from about 400 meters ² / g to about 1200 m ² / g. Pore volume of the catalyst is at least about 0.3 cm ³ / g, preferably ranging from about 0.4 cm ³ / g to about 2.2 cm ³ / g.

  There are a number of methods for producing the catalyst support TUD-1, but these methods can be classified into two types, depending on the starting inorganic oxide: (1) an organic-containing precursor, and (2) Inorganic precursor. In the former case, the inorganic oxide precursor is an alkoxide having a desired element selected from silicon, aluminum, titanium, vanadium, zirconium, gallium, manganese, zinc, chromium, molybdenum, nickel, cobalt, and iron, such as tetraethyl An organic silicate such as orthosilicate (TEOS) or an organic source of aluminum oxide such as aluminum isopropoxide may be preferred. TEOS and aluminum isopropoxide are commercially available from well-known suppliers.

  It is preferred that the pH of the solution is maintained above 7.0. In some cases, the aqueous solution may contain other metal ions, such as the metal ions described above. After stirring, an organic mesopore former that binds to silica (or other inorganic oxide) species by hydrogen bonding is added to the aqueous solution and mixed. Organic pore formers are glycols (compounds containing two or more hydroxy groups) such as glycerol, diethylene glycol, triethylene glycol, tetraethylene glycol, propylene glycol and the like, or triethanolamine, sulfolane , Tetraethylenepentamine, and diethyl glycol dibenzoate are preferably selected from the group consisting of one or more. The organic pore forming agent needs to have low hydrophobicity so as not to form a separated phase in the aqueous solution, and is preferably added by dropwise addition to the aqueous inorganic oxide solution with stirring. After time has elapsed (eg, about 1-2 hours), the mixture forms a high viscosity gel. It is preferable to stir the mixture during this period to facilitate mixing of the components. The mixture preferably contains an alkanol. The alkanol can be added to the mixture and / or can be formed in-situ by decomposition of the inorganic oxide precursor. For example, TEOS produces ethanol when heated. Propanol can be produced by decomposition of aluminum isopropoxide.

  A second type of synthetic route to obtain the same gel is to use an inorganic precursor as a starting material. Preferred inorganic precursors are oxides and / or hydroxides with the desired elements selected from silicon, aluminum, titanium, vanadium, zirconium, gallium, manganese, zinc, chromium, molybdenum, nickel, cobalt, and iron including. The precursor is first mixed with one or more pore formers and for a predetermined time (eg, 2-10 hours), 120-250 for sufficient time for the inorganic precursor to convert to an organic-containing composite. Heated to ° C. The composite is then mixed with water and hydrolyzed to obtain a homogeneous high viscosity gel.

  The gel obtained by the above two methods is then placed at about 5 ° C. to about 45 ° C., preferably room temperature (aging) to complete the hydrolysis and polycondensation of the inorganic oxide source. Aging is preferably carried out for up to about 48 hours, generally about 2 to 30 hours, more preferably about 10 to 20 hours. After the aging step, the gel is heated in air at about 90 ° C. to 100 ° C. for a time sufficient to remove moisture from the gel (eg, about 6 to about 24 hours). The organic pore former that assists in the formation of mesopores preferably remains in the gel during the drying stage. Accordingly, preferred organic pore formers have a boiling point of at least about 150 ° C.

  The dried material containing the organic pore former is heated at a temperature at which a significant number of mesopores are formed. This pore-forming step is performed at a temperature that exceeds the boiling point of water and reaches the boiling point of the organic pore-forming agent. Generally, mesopore formation is performed at a temperature of about 100 ° C to about 250 ° C, preferably about 150 ° C to about 200 ° C. The hole forming step may optionally be performed hydrothermally under autogenous pressure in a sealed container. The size and volume of the final product mesopores are affected by the duration and temperature of the hydrothermal step. In general, increasing the processing temperature and duration increases the proportion of mesopore volume of the final product.

  After the pore formation step, the catalyst material is calcined at a temperature of about 300 ° C. to about 1000 ° C., preferably about 400 ° C. to about 700 ° C., more preferably about 500 ° C. to about 600 ° C. for a period of time sufficient to calcine the material Maintain at the firing temperature. The duration of the firing step depends in part on the firing temperature, but is usually about 2 hours to about 40 hours, preferably 5 hours to 15 hours.

  In order to prevent heat spots, the temperature needs to be gradually raised. Preferably, the temperature of the catalyst material is from about 0.1 ° C./min to about 25 ° C./min, more preferably from about 0.5 ° C./min to about 15 ° C./min, most preferably from about 1 ° C./min to about It is necessary to increase to the firing temperature at a rate of 5 ° C./min.

  During calcination, the structure of the catalyst material is finally formed while organic molecules are excluded from the material and decomposed.

  The calcination method for removing the organic pore forming agent can be replaced by extraction using an organic solvent such as ethanol. In this case, the pore-forming agent can be recovered and reused.

  Similarly, the catalyst powder of the present invention may be mixed with a binder such as silica and / or alumina and then formed into a desired shape (eg, pellets, rings, etc.) by injection or other suitable method. .

  The catalyst includes at least one metal component selected from Group VIII of the Periodic Table of Elements. This includes iron, cobalt, nickel, and noble metals, ie, platinum, palladium, ruthenium, rhodium, osmium, and iridium. Particularly preferred metals include platinum, palladium, rhodium, iridium, and nickel. The amount of Group VIII metal is at least about 0.1% by weight, based on the total weight of the catalyst.

  The Group VIII metal can be incorporated into the inorganic oxide by any suitable method. For example, ion exchange or a solution of a soluble and decomposable Group VIII metal compound is impregnated in an inorganic oxide, and then the impregnated inorganic oxide is washed, dried, and fired, for example, The process includes decomposing the Group VIII metal compound to produce an activated catalyst having a free Group VIII metal in the pores of the inorganic oxide. Suitable Group VIII metal compounds include salts such as nitrates, chlorides, ammonium complexes, and the like.

  The inorganic oxide catalyst impregnated with the Group VIII metal is optionally washed with water to remove some anions. Drying the catalyst to remove water and / or other volatile compounds can be performed by heating the catalyst to a drying temperature of about 50 ° C to about 190 ° C. Calcination for activating the catalyst can be performed at a temperature of about 150 ° C. to about 600 ° C. for a sufficient time. In general, calcination can be carried out for 2 to 40 hours, depending at least in part on the calcination temperature.

  Optionally, one or more zeolites may be incorporated into the catalyst and dispersed throughout the mesoporous matrix. Zeolite is preferably added to the inorganic oxide precursor-water solution prior to formation of the mesoporous structure. Suitable zeolites include, for example, FAU, EMT, BEA, VPI, AET, and / or CLO. The zeolite is preferably present in an amount of 0.05% to 50% by weight, based on the total weight of the catalyst.

  Another preferred hydrogenation type includes the selective removal of impurities in the feedstock containing hydrocarbons. More specifically, selective hydrogenation of a compound having a triple bond and / or a compound having two or more double bonds relative to a compound having one double bond, and one two double bonds. It relates to the selective hydrogenation step of a compound having two adjacent double bonds with respect to the compound separated by the above single bond.

  Such reactions include, but are not limited to, selective hydrogenation of acetylene and / or diene impurities in the feed containing at least one monoolefin. Further examples include selective hydrogenation of acetylene in an ethylene stream, selective hydrogenation of methylacetylene and propadiene in a propylene stream, selective hydrogenation of butadiene in a butene stream, and a feed containing 1,3-butadiene. Including selective hydrogenation of vinyl and ethylacetylene, and 1,2-butadiene.

  In the petrochemical industry, the product stream contains one or more monoolefins and acetylene and / or diene compounds as impurities. Acetylene impurities include acetylene, methylacetylene, and diacetylene. Diene impurities include 1,2-butadiene, 1,3-butadiene, and propadiene.

  Such product streams typically undergo selective hydrogenation to minimize / remove acetylene and / or diene impurities without hydrogenating the desired monoolefin. This process can be carried out by selective catalytic hydrogenation using a catalyst.

  This catalyst comprises a metal, preferably a noble metal, and optionally a binder supported on the porous material of the present invention. The catalyst may contain additional metals that are used as promoters.

  Selective hydrogenation of acetylene and / or diene impurities is carried out by one-stage hydrogenation in the presence of the catalyst described above. The raw material is introduced as a liquid and partly or completely vaporized during the hydrogenation. The feed is selectively hydrogenated and a hydrogen gas stream is introduced into the reactor at a temperature of about 0 ° C to 50 ° C. The reactor is operated at a pressure ranging from about 1379 kPa to about 3448 kPa (200 psi to 500 psi). Depending on the level of acetylene and / or diene impurities in the feed, inlet temperature, and acceptable outlet temperature, a portion of the product may need to be recycled to the reaction zone.

  The amount of hydrogen introduced into the reactor is based on the amount of impurities in the feedstock. Hydrogen can be introduced into the reactor along with a suitable diluent such as methane.

  Appropriate liquid space velocities need to be used, as will be apparent to those skilled in the art.

  The following examples illustrate the features of the present invention.

<Example 1>
In this example, a method for synthesizing Si-TUD-1 using silicon alkoxide as the silica source is described. 736 parts by mass of tetraethylorthosilicate (98%, manufactured by ACROS) was mixed with 540 parts of triethanolamine (TEA) (97%, manufactured by ACROS) while stirring. After 30 minutes, 590 parts of water was slowly added to the above mixture with stirring. After an additional 30 minutes, 145 parts of tetraethylammonium hydroxide (TEOH) (35% by weight) was added to the above mixture to obtain a homogeneous gel. The gel was left at room temperature for 24 hours. The gel was then dried at about 98 ° C. for 18 hours and baked at 600 ° C. for 10 hours in air at a heating rate of 1 ° C./min.

The X-ray diffraction (XRD) pattern of the final material showed a strong 2Θ peak of less than 2 ° indicating a mesoporous structure. As a result of the BET measurement by nitrogen absorption, the surface area was 683 m ² / g, the average pore diameter was about 4.0 nm, and the total pore volume was about 0.7 cm ³ / g.

<Example 2>
In this example, a method for synthesizing Si-TUD-1 using silica gel as a silica source will be described. First, 24 parts of silica gel, 76 parts of TEA, and 62 parts of ethylene glycol (EG) were placed in a reactor equipped with a condenser. After the contents in the reactor were mixed well with a mechanical stirrer, the mixture was heated to 200-210 ° C. with stirring. This removed most of the water produced during the reaction from the top of the condenser along with some of the EG. On the other hand, most of EG and TEA remained in the reaction mixture. After about 8 hours, the heating was stopped, and after cooling to room temperature, a light brown paste-like composite mixture was recovered.

  Next, 100 parts of water was added to 125 parts of the composite liquid obtained above under stirring conditions. After stirring for 1 hour, the mixture formed a high viscosity gel that was left at room temperature for 2 days.

  Next, the high-viscosity gel was dried at 98 ° C. for 23 hours, then placed in an autoclave and heated to 180 ° C. over 6 hours. Finally, it was fired at 600 ° C. for 10 hours in air at a heating rate of 1 ° C./min.

The X-ray diffraction (XRD) pattern of the final material showed a strong 2Θ peak below 2 ° indicating a mesoporous structure. As a result of BET measurement by nitrogen absorption, the surface area was 556 m ² / g, the average pore diameter was about 8.1 nm, and the total pore volume was about 0.92 cm ³ / g.

<Example 3>
In this example, the synthesis of Al-Si-TUD-1 is described. First, 250 parts of silica gel, 697 parts of TEA, and 287 parts of ethylene glycol (EG) were placed in a reactor equipped with a condenser. After the contents in the reactor were mixed well with a mechanical stirrer, the mixture was heated to 200-210 ° C. with stirring. This removed most of the water produced during the reaction from the top of the condenser along with some of the EG. On the other hand, most of EG and TEA remained in the reaction mixture. After about 3 hours, the reactor was cooled to 100 ° C. and another mixture containing 237 parts of aluminum hydroxide, 207 g of EG, and 500 g of TEA was added to the reactor. The mixture was heated again to 200-210 ° C. and after 4 hours the heating was stopped. After the mixture was cooled to room temperature, a light brown pasty composite liquid was recovered.

  Next, 760 parts of water and 350 parts of tetraethylammonium hydroxide were added to the composite liquid obtained above under stirring conditions. After stirring for 1 hour, the mixture formed a high viscosity gel that was left at room temperature for 1 day.

  Next, the high viscosity gel was dried at 98 ° C. for 23 hours, then placed in an autoclave and heated to 180 ° C. over 16 hours. Finally, it was fired at 600 ° C. for 10 hours in air at a heating rate of 1 ° C./min.

The X-ray diffraction (XRD) pattern of the final material showed a strong 2Θ peak below 2 ° indicating a mesoporous structure. As a result of BET measurement by nitrogen absorption, the surface area was 606 m ² / g, the average pore diameter was about 6.0 nm, and the total pore volume was about 0.78 cm ³ / g.

<Example 4>
This example illustrates the preparation of a 0.90 wt% iridium / Si-TUD-1 catalyst by the incipient wetness method. 0.134 parts of indium chloride (IU) was dissolved in 5.2 parts of deionized water. This solution was added to 8 parts of Si-TUD-1 obtained in Example 1 with mixing. The powder was dried at 25 ° C.

  The powder was then reduced in a stream of hydrogen at 100 ° C. for 1 hour for dispersion measurement using CO chemisorption, then heated to 350 ° C. at 5 ° C./min and held at this temperature for 2 hours. CO chemisorption shows a 75% dispersion of the metal and an Ir: CO stoichiometric ratio of 1 is estimated.

<Example 5>
This example describes the preparation of 0.9 wt% palladium and 0.3 wt% platinum / Si-TUD-1 by the incipient wetness method. First, Al-Si-TUD-1 obtained in Example 3 was extruded. Next, 70 parts of 1/16 "extrudate were impregnated with 0.42 parts of an aqueous solution of tetraamine platinum nitrate, 12.5 parts of an aqueous solution of tetraamine palladium nitrate (5% Pd), and 43 parts of water. Impregnated Al- Si-TUD-1 was left at room temperature for 6 hours prior to drying for 2 hours at 90 ° C. Finally, the dried material was calcined in air at 350 ° C. for 4 hours at a heating rate of 1 ° C./min. CO chemisorption was used to measure noble metal dispersion, and then the powder was reduced in a hydrogen stream at 100 ° C. for 1 hour, then heated to 350 ° C. at 5 ° C./min and held at this temperature for 2 hours. A 51% dispersion of the metal is measured and a Pt: CO stoichiometry of 1 is estimated.

<Example 6>
This example describes the production of 0.46 wt% platinum / Si-TUD-1 by the incipient wetness method. 0.046 parts of tetraamine platinum nitrate (II) was dissolved in 4.1 parts of deionized water. This solution was added to 5 parts of Si-TUD-1 obtained in Example 1 with mixing. The powder was dried at 25 ° C.

  The powder was then reduced in a stream of hydrogen at 100 ° C. for 1 hour for dispersion measurement using CO chemisorption, then heated to 350 ° C. at 5 ° C./min and held at this temperature for 2 hours. A 72% variance of the sample is measured and a Pt: CO stoichiometry of 1 is estimated.

<Example 7>
21 parts of Si-TUD-1 obtained in Example 1 was suspended in deionized water. Nitric acid was added to adjust the pH of the solution to 2.5. This exchange was performed for 5 hours. The solution was then drained. The Si-TUD-1 was then washed 5 times with deionized water. This Si-TUD-1 was then placed in 600 parts of deionized water. The pH of this solution was adjusted to 9.5 using ammonium nitrate. This exchange was performed for 1 hour. During this exchange, ammonium nitrate was added as needed to maintain the pH at 9.5. After the exchange, Si-TUD-1 was washed 5 times with deionized water. Next, Si-TUD-1 was dried at 25 ° C. Using this acid / base-treated Si-TUD-1, 0.50% palladium / Si-TUD-1 was produced by an indirect wetness of tetraamine platinum nitrate (II). 0.071 parts of palladium salt was dissolved in 4.1 parts of deionized water. This solution was added to 5 parts Si-TUD-1 with mixing. The powder was dried at 25 ° C. Next, the catalyst powder was calcined at 350 ° C. for 2 hours in air at a heating rate of 1 ° C./min.

  The powder was then reduced in a stream of hydrogen at 100 ° C. for 1 hour for dispersion measurement using CO chemisorption, then heated to 350 ° C. at 5 ° C./min and held at this temperature for 2 hours. A 96% dispersion of the sample is measured and a Pd: CO stoichiometry of 1 is estimated

<Example 8>
This example illustrates the production of 0.25% palladium / Si-TUD-1 by the incipient wetness method of tetraamine palladium nitrate (II) using acid / base treated TUD-1 (Example 7). 0.035 parts of palladium salt was dissolved in 3.9 parts of deionized water. This solution was added to 5 parts Si-TUD-1 with mixing. The powder was dried at 25 ° C. Next, the catalyst powder was calcined at 350 ° C. for 2 hours in air at a heating rate of 1 ° C./min.

  CO chemisorption was used to measure precious metal dispersion, then the powder was reduced in a hydrogen stream at 100 ° C. for 1 hour, then heated to 350 ° C. at 5 ° C./min and held at this temperature for 2 hours. 90% dispersion of the sample is measured and a Pt: CO stoichiometry of 1 is estimated

<Example 9>
A 0.38 wt% palladium / 0.23% wt platinum / Si-TUD-1 catalyst was prepared as follows. Acid / base treated Si-TUD-1 (Example 7) was used to produce 0.38% palladium TUD-1 by an incipient wetness of tetraamine palladium nitrate (II). 0.053 parts of palladium salt was dissolved in 3.75 parts of deionized water. This solution was added to 5 parts of Si-TUD-1 with mixing. The powder was dried at 25 ° C. Next, the catalyst powder was calcined at 350 ° C. for 2 hours in air at a heating rate of 1 ° C./min.

  A 0.23% wt platinum impregnation on the catalyst was prepared by an incipient wetness of tetraamine platinum (II) nitrate. 0.018 part of platinum salt was dissolved in 3.25 parts of deionized water. This solution was added to 4.02 parts of 0.38 wt% Pd / Si-TUD-1 with stirring. The powder was dried at 25 ° C.

  Next, for dispersion measurements using CO chemisorption, the powder was reduced in a hydrogen stream at 100 ° C. for 1 hour, then heated to 350 ° C. at 5 ° C./min and held at this temperature for 2 hours. 81% variance of the sample is measured and a stoichiometric ratio of 1 for Pd: CO and Pt: CO is estimated

<Example 10>
A 0.19 wt% palladium / 0.11 wt% platinum / Si-TUD-1 catalyst was prepared as follows. 0.19 wt% palladium / Si-TUD-1 was prepared by an indirect wetness of tetraamine palladium nitrate (II) using acid / base treated Si-TUD-1 (Example 7). 0.027 parts of the palladium salt was dissolved in 3.5 parts of deionized water. This solution was added to 5 parts of Si-TUD-1 with mixing. The powder was dried at 25 ° C. The catalyst powder was then calcined at 350 ° C. for 2 hours in air at a heating rate of 1 ° C./min.

  A 0.11 wt% platinum impregnation on the catalyst was prepared by an incipient wetness of tetraamine platinum (II) nitrate. 0.009 part of platinum salt was dissolved in 3.27 parts of deionized water. This solution was added to 4.05 parts of 0.19% Pd / Si-TUD-1 with mixing. The powder was dried at 25 ° C.

  The powder was then reduced in a stream of hydrogen at 100 ° C. for 1 hour for dispersion measurement using CO chemisorption, then heated to 350 ° C. at 5 ° C./min and held at this temperature for 2 hours. The dispersion of 54% of the sample is measured and a stoichiometric ratio of 1 for Pd: CO and Pt: CO is estimated.

<Example 11>
The TUD-1 catalyst was evaluated in a 1 "reactor with actual continuous feed and compared to a commercial catalyst. Table 1 summarizes the processing conditions. Table 2 shows the feedstock and effluent characteristics and final production. The TUD-1 catalyst gives a final product with only 5% aromatics and the commercial catalyst gives a product with 10% aromatics at high space velocities. The TUD-1 catalyst showed higher aromatic saturation activity.

<Example 12>
In this example, an aluminum-based TUD-1 was produced. 65 parts by weight of isopropanol and 85 parts of ethanol were added to a container containing 53 parts of aluminum isopropoxide. After stirring at 50 ° C. for about 4 hours, 50 parts of tetraethylene glycol (TEG) was added dropwise with stirring. After further stirring for 4 hours, 10 parts of water was added with stirring together with 20 parts of isopropanol and 18 parts of ethanol. After stirring for 30 minutes, the mixture became a white suspension, which was then left at room temperature for 48 hours, followed by drying in air at 70 ° C. for 20 hours to obtain a solid gel. This solid gel was heated in an autoclave at 160 ° C. for 2.5 hours, and finally calcined in air at 600 ° C. for 6 hours to produce mesoporous aluminum oxide.

In the XRD pattern of the obtained calcined mesoporous aluminum oxide, there was a strong 1.6 ° 2Θ peak characteristic of mesostructured materials. As a result of the N ₂ porosimeter, the pore size distribution was concentrated around 4.6 nm. As a result of ²⁷ Al NMR spectroscopy, three peaks corresponding to 4, 5 and 6 coordinated aluminum were observed at 75, 35 and 0 ppm, respectively. That is, this was a typical mesoporous material of the present invention with 4, 5 and 6 coordinated aluminum.

<Example 13>
This example illustrates the use of the composition of the present invention as a catalyst support for hydrogenation. First, Al-TUD-13.13 parts of Example 12 (“Sample 12”) were impregnated with 2 parts of a 3.1 wt% Pt (NH ₃ ) ₄ (NO ₃ ) ₂ aqueous solution by the incipient wetness method. . After drying and calcination in air at 350 ° C. for 2 hours, 1250 parts of the impregnated sample were charged into the reactor and then reduced with hydrogen at 300 ° C. for 2 hours.

As a probe reaction, hydrogenation of mesitylene was carried out in a fixed bed reactor having a feed of 2.2 mol% mesitylene concentration in hydrogen under a total pressure of 600 kPa (6 bar). In order to determine the rate constant of the catalyst, the reaction temperature was varied in the range of 100-130 ° C in 10 ° C increments. The corrected contact time based on catalyst mass was 0.6 g _cat. * Maintained constant at min * l ^-1 . The first-order reaction rate constant based on the catalyst mass was 0.15 g _cat. ⁻¹ * min− ¹ * 1.

<Example 14>
This example illustrates the selective hydrogenation of acetylene and diene. Pd-AgAl-TUD-1 catalyst was prepared in the form of 1/16 "extrudate and ground to 24/36 mesh particles for laboratory performance testing. In a 0.75" OD tubular reactor, Selective hydrogenation was performed. The feedstock consisted of 0.8% methylacetylene, 0.3% propadiene, 22% propylene, and a balance amount of isobutane. Hydrogen was dissolved in the hydrocarbon stream. The molar ratio of hydrogen / (methylacetylene + propadiene) was about 0.75. This mixture was then transferred to another reactor. LHSV was maintained at about 367. At the end of the reaction, conversion and selectivity were measured. Selectivity was defined as the propylene produced / [converted (methylacetylene + propadiene)] × 100. At 49 ° C. and about 2758 kPaG (400 psig), the (methylacetylene + propadiene) conversion was 29% and the selectivity was 71%.

  While the foregoing description includes numerous details, these details are not intended to limit the scope of the invention and are to be construed as merely illustrative of preferred embodiments of the invention. Those skilled in the art will envision many other possibilities within the scope and spirit of the invention as defined by the claims appended hereto.

Claims (28)

A method for hydrogenating a hydrocarbon feedstock containing unsaturated components, comprising:
a) having at least 97 volume percent interconnected mesopores based on mesopores and micropores and having a BET surface area of at least 300 m ² / g and a pore volume of at least 0.3 cm ³ / g Providing a catalyst comprising at least one Group VIII metal on an amorphous mesoporous inorganic oxide support having:
b) contacting the hydrocarbon feed with hydrogen in the hydrogenation reaction zone under hydrogenation reaction conditions in the presence of the catalyst to obtain a product having a reduced content of unsaturated components. Method.   The method of claim 1, wherein the Group VIII metal is a noble metal.   The method of claim 2, wherein the noble metal is selected from the group consisting of palladium, platinum, rhodium, ruthenium, and iridium.   The method of claim 1, wherein the Group VIII metal is nickel.   The method of claim 1, wherein the Group VIII metal has a percentage composition of at least about 0.1 weight percent, based on the total weight of the catalyst. The mesoporous inorganic oxide support has a BET surface area of about 400 meters ² / g to about 1,200m ² / g, and a pore volume of about 0.4 cm ³ / g to about 2.2 cm ³ / g, wherein Item 2. The method according to Item 1.   The method of claim 1, wherein the unsaturated component of the hydrocarbon feedstock comprises an aromatic compound and / or an olefin.   The method of claim 7, wherein the method is dearomatization. The hydrogenation reaction conditions are from about 0.99 ° C. to about 400 ° C. of temperature, hydrogen partial pressure of about 1379kPa~ about 13790KPa (about 200psi~ about 2,000 psi), about 0.2 hr ^-1 to about 10.0 hours ^{- 2. The process of claim 1 comprising} an LHSV of ¹ and a hydrogen cycling ratio of about 500 SCF / Bbl to about 20,000 SCF / Bbl. The hydrogenation reaction conditions include a temperature of about 260 ° C. to about 650 ° C., a hydrogen partial pressure of about 3448 kPa to about 10343 kPa (about 500 psi to about 1,500 psi), about 0.5 hours ⁻¹ to about 3.0 hours ^{− The process of claim 1 comprising} an LHSV of ¹ and a hydrogen cycling ratio of from about 2,000 SCF / Bbl to about 15,000 SCF / Bbl. The hydrogenation reaction conditions are a temperature of about 275 ° C. to about 330 ° C., a hydrogen partial pressure of about 4137kPa~ about 8274KPa (about 600psi~ about 1,200 psi), from about 1.0 hr ^-1 to about 2.0 hr ^{- 2.} The method of claim 1 comprising an LHSV of ¹ and a hydrogen cycling ratio of from about 3,000 SCF / Bbl to about 13,000 SCF / Bbl.   The method of claim 1, wherein the catalyst further comprises a zeolite.   The method of claim 12, wherein the zeolite is selected from the group consisting of FAU, EMT, BEA, VFI, AET, CLO, and combinations thereof.   The method of claim 12, wherein the amount of zeolite is from about 0.05% to about 50.0% by weight, based on the total weight of the catalyst.   The method of claim 1, wherein the hydrocarbon feedstock comprises a lubricant base stock.   The method of claim 1, wherein the hydrocarbon feedstock contains more than about 70% by weight aromatic compounds.   The method of claim 1, wherein the hydrocarbon feedstock contains more than about 50% by weight aromatic compounds.   The method of claim 1, wherein the hydrogenation reaction zone comprises at least one catalyst fixed bed.   The method of claim 18, wherein the hydrogenation reaction zone includes at least first and second spaced apart catalyst fixed layers, wherein the first fixed layer effluent is directed into the second fixed layer. .   20. The method of claim 19, further comprising the step of cooling the effluent before the first fixed bed effluent enters the second fixed bed.   The method of claim 1, further comprising heating the feedstock to a reaction temperature in a furnace after preheating the feedstock in a feed / effluent heat exchanger.   The method of claim 1, wherein the method comprises dearomatization of a hydrocarbon feedstock containing an aromatic compound.   The method of claim 1, wherein the process comprises hydrogenating a hydrocarbon lubricant basestock feedstock having a bromine number greater than 5 in the presence of superatmospheric hydrogen to produce a lubricant product having a bromine number less than 3. the method of.   The process of claim 1, wherein the process comprises selective hydrogenation of acetylene and / or diene impurities in a feedstock containing at least one monoolefin.   The method of claim 1, wherein the method comprises selective hydrogenation of olefin and / or diene impurities in a feedstock containing at least one aromatic compound.   The compound in which the impurity includes acetylene and one or more compounds having adjacent double bonds, and the hydrocarbon raw material has two or more double bonds separated by at least one single bond 25. The method of claim 24, comprising: A method of stabilizing a lubricating oil containing unsaturated components,
(A) hydrocracking a hydrocarbon raw material of lubricant viscosity in a hydrocracking zone to obtain an effluent;
(B) at least a portion of the effluent of the hydrocracking zone is connected to at least 97 volume percent of meso fines in a catalytic hydrogenation zone under superatmospheric hydrogen pressure, based on mesopores and micropores. At least one Group VIII metal is included on an amorphous mesoporous inorganic oxide support having pores and having a BET surface area of at least 300 m ² / g and a pore volume of at least 0.4 cm ³ / g And catalytically hydrogenating at least a portion of the hydrocracking zone effluent by contacting with a catalyst to obtain a lubricant product having a reduced content of unsaturated components. A method of stabilizing a lubricating oil containing unsaturated components,
(A) hydrocracking a hydrocarbon raw material of lubricant viscosity in a hydrocracking zone to obtain an effluent;
(B) at least a portion of the effluent of the hydrocracking zone is connected to at least 97 volume percent of meso fines in a catalytic hydrogenation zone under superatmospheric hydrogen pressure, based on mesopores and micropores. Contacting with a catalyst comprising nickel on an amorphous mesoporous inorganic oxide support having pores and having a BET surface area of at least 300 m ² / g and a pore volume of at least 0.4 cm ³ / g, Catalytically hydrogenating at least a portion of the hydrocracking zone effluent to obtain a lubricant product having a reduced content of unsaturated components.